Preparation of high octane number motor fuel blending stocks



Jan. 17, 1961 E.s. coDDou, JR, ETAL 2,968,605

PREPARATION OF HIGH OCTANEZ NUMBER MOTOR FUEL BLENDING STOCKS Filed July 28, 1958 Nm m V @E a @Jr vvvvnxl Rd n, Y y memm mm m M m m M mm EBEE vm N o n n H summum mx wzzsm mm m5. m uw m wzomor o tct MF. J v MIN W f d A a mm v mw a 0 wkvv Mv vvvv Av W @n LN vv v 4 Sm um @N mm um @mi w S mm G .vvvv @zuo Q ;Q u \1e m/ h J zQzoCE wmmwwwm MM, @issu zQzmo Zimio Essi mm .V Atm, wm, Il H M|.\ mmmw .Nm 053mb 5.5i v v .Gg QQ Zmwm United States Patent' PREPARATION OF HIGH OCTANE NUMBER MOTOR FUEL BLENDING STOCKS Eugene S. Coddou, Jr., Houston, and Walker F. Johnston, Jr., La Marque, Tex., and David J. Bellman, Ponce, Puerto Rico, assignors to The American Oil Company, Texas City, Tex., a corporation of Texas i Filed July 28, 1958, Ser. No. 751,469

5 Claims. (Cl. 208-70) This invention relates to the production of high octane number motor fuel blending stocks, particularly the production of 98-100 research (CFR-R) clear octane number blending stocks by a process integrating the favorable advantages of catalytic cracking and catalytic hydroforming.

In recent years the steadily increasing octane level for premium grade gasolines, i.e., gasolines having octane numbers of 97-98 research and higher, have imposed a severe burden on the petroleum refining industry. The problem is basically one of economics, since modern technology affords processes which, theoretically at least, are capable of producing limited amounts of gasolines having an octane number substantially in excess of any foreseeable engine requirement. However, when attempts are made to apply theory to practice, it is frequently found in this art that an economic burden is incurred which militates against installing processes which at iirst blush would appear attractive. Therefore, the art is faced with the problem of meeting market requirements for premium grade gasolines which can be produced in large volumes by means of conventional refining processes, and yet accomplish this production in an economic manner. lt is therefore a primary object of the instant invention to provide an economically attractive process for producing high octane motor fuel blending stocks by means of a low cost process, integrating the operation of conventional petroleum relining units in a manner such that operating conditions of each individual unit are correlated so as to provide a total blending stock yield which is substantially in excess of yields of similar quality stocks hitherto obtainable.

In accordance with the practice of the present invention, the process described herein includes the catalytic cracking of petroleum gas oils at defined low severity conditions to produce a relatively high yield of low octane gasoline, about 84 to 94 octane number clear on a debutanized basis. The catalytically cracked reaction products are fractionally distilled to separate a relatively high octane number light gasoline from a heavy gasoline boiling chiey within the range of about 200 to about 400 F., next the latter is subjected to catalytic hydrogenation to selectively hydrogenate olefinic constituents. Following olefin hydrogenation, and preferably following a stripper or scrubber to remove acidic hydrogenated contaminants in the heavy gasoline, the gasoline is subjected to catalytic hydroforming to produce a 98-108 research clear octane number blending stock. A portion of the net hydrogen from the hydroforming operation is supplied to the olefin hydrogenation step wherein said gas is req uired for the reaction. Thus, by conducting catalytic ICC cracking at low severity so as to obtain relatively high yields of low octane number gasoline, and subjecting the lowest octane fraction of this gasoline to hydrogenation and hydroforming, an extremely high yield of blending stock is obtained which benefits from a coordinated operation of the cracking unit to give high yields, with high severity operation of the catalytic reforming unit to provide high octane number characteristics.

The invention will be more fully understood by reference to the attached drawing showing a schematic flow sheet of the process of the present invention. Auxiliaries and utilities such as pumps, meters, gages, steam and electric power lines, and the like have been omitted for reasons of clarity and simplicity, and those skilled in the art will readily appreciate the location and type of these details.

Operations commence by charging 50,000 barrels per day of gas oil feedstock to the catalytic cracking unit via line 1. This gas oil preferably comprises or consists of a refinery crude which has been previouslyl subjected to atmospheric or low pressure distillation to remove virgin fractions boiling at least in the gasoline boiling range of up to 400 F., and optionally for the removal of certain heavier fractions for use as kerosene and heater oils. The gas oil has been vacuum distilled to remove a high boiling residuum; the feedstock may contain fractions derived from thermal cracking or viscosity breaking operations. Feedstock line 1 connects with the catalytic cracking unit at junction 2, where a descending stream of freshly regenerated fluidized particulate cracking catalyst obtained from regenerator 13 by means of standpipe 20 is mixed with the feedstock. Prior to mixing, the feedstock may be at a temperature ranging from about 50 F. to about 650 F. if preheating is employed, and after mixing with the hot regenerated catalyst the feedstock is vaporzed and, at a temperature preferably within the range of about 750 F. to about 900 F. ascends from riser line 3 into the bottom of catalytic cracking reactor 4. Vaporized feedstock and cracking catalyst enter a space 5 at the bottom of reactor 4 and is distributed into dense phase region 7 by means of perforated grid 6.. The cracking catalyst forms a relatively dense uidized bed 7 which behaves much as a common liquid, while the Vaporized feedstock is cracked by contact with the cracking catalyst into lower boiling constituents boiling predominantly in the ISO-400 F. gasoline boiling range. The vaporzed reaction products ascend through dense phase 7 and into a dilute phase or disengaging space at the top of reactor 4, thereafter passing through cyclone separator 8 to product transfer line 9. ln cyclone separator 8, it is shown symbolically as a single cyclone but may be a plurality of parallel or serially connected cyclones of suitable size. Entrained catalyst is separated from the vaporzed reaction products and returned to dense phase by dipleg means as shown. Catalyst is withdrawn from reactor 4 by passing through perforated grid 10 into a stripper section within reactor 4 wherein a small quantity of steam admitted through line 11 strips occluded oil vapors from the catalyst particles. The stripped particles then descend via standpipe 12 and are cycled to regenerator 13.

In regenerator 13, the catalyst forms a liuidized dense bed atthe lower portion of the regenerator chamber, and carbonaceous deposits on the catalyst are burned oli at a temperature generally within the range of about 900- 1,200 F., preferably 1,0001,050 F., by the admission of air through line 15, plenum chamber 16 and distribution grid 17. Air ascends through dense phase 14 and thence into the dilute phase region and into cyclone separator 18 (also shown symbolically as one cyclone) where entrained catalyst is separated and returned to dense phase 14 while the spent or ue gas is released through stack 19. After regeneration, the catalyst is in condition for cycling to reactor 4 via regenerated catalyst standpipe 20.

An important feature in the practice of the instant invention is regulation of the catalytic cracking operation to provide a relatively high yield of low octane number gasoline. Ordinarily, with cracking catalysts such as the well-known silica-alumina synthetic microspheroidal catalyst an optimum cracking temperature of between about 900 F. and 950 F. is employed so as to produce a gasoline having a clear research octane number, on a debutanized basis, from about 93 to 96. The temperature, as is well known, is one of several variables and other cracking catalysts, typified by silica-magnesia and natural clay or Filtrol-type catalyst may require a higher temperature for equivalent octane number. The yield yof debutanized gasoline will vary greatly depending upon the charge stock used, but when using the same charge stock, reducing the temperature by about 150 F. will increase the yield in the order of 5%. Therefore, in accordance with the present invention the reactor ternperature and other variables such as nature of catalyst, catalyst to oil ratio, and catalyst inventory in reactor 4 are controlled in known manner to yield a product having a debutanized gasoline of only 84-94 octane number. Thus, while the octane number is lower, the yield is higher than in conventional operations, where high octane-lower yield catalytically cracked gasoline is blended without further processing.` On the other hand, the present inventiony contemplates splitting the higher' yield of debutaniiz'ed catalytically'cracked gasoline into a high octane-low boiling fraction and a low octane-high boiling fraction. Further processing the high boiling fraction by hydrogenation and high severity hydroforming renders the catalytic cracking operations which tend to produce more gasoline but of lower octane number, a desirable modification.

The vaporized reaction mixture leaving reactor 4 via transfer line 9 is conducted to main fractionating tower 22 where the components are roughly separated by means of differences in their boiling points. Beginning with the heaviest fraction, a high boiling catalyst slurry oil `is withdrawn through line 33 and passed through cooling means 34 and thereafter divided into two streams, one of them 35 being a net slurry oil product containing less than about one percent-suspended and entrained catalyst, and another portion being returned through line 36 to aid in cooling the reaction products. Because this slurry oil in the bottom of main fractionator 22 is quite hot, cooling means 34 desirably constitutes the reboilers of distillation towers throughout the process herein described. Higher up in fractionator 22 a heavy cycle oil product is obtained from draw-Gif tray 31 and collected through line 32; this heavy cycle oil may be conducted to a stripping tower where steam is supplied at the bottom so as to strip low boiling components of the cycle oil and return, along with thesteam, back to main fractionator 22. Similarly a light cycle oil product is collected by way of draw-off 29 and line 30. The slurry oil, the heavy oil, and the light cycle oil may be cycled in whole or in part to juncture 2 and thence back into catalytic cracking reactor 4 for further exposure to the cracking catalyst or may be collected as final products. Slurry oil may be used in amounts ranging from about 10 to 50% as cutter stock for refinery residual fuels, while heavy cycle and light cycle oil may be used as such or in blends as industrial and home heating fuels.

All components of the cracking unit products boiling below about- 400 F. (the temperature may range from 350 F. to 450 F. depending on local reiinery gasoline end-point specifications) are taken overhead through vapor line 23 and main fractionator 22. Condenser 24 condenses a major portion `of the butanes and heavier components, along with the stripping steam, and the con densate along with uncondensed gaseous materials are charged to overhead receiver 25. A portion of the liquid hydrocarbons are returned to fractionator 22 via line 26 to serve as reflux, while the balance is pumped to splitter fractionator 37. Uncondensed reaction products are withdrawn from the vapor space of overhead receiver 25 and transmitted through line 28 to light ends recovery unit, not shown, which may employ at least a portion of the catalytically cracked and unstabilized gasoline (obtained via line 27 from receiver 25) or a higher boiling material such as light cycle oil to scrub out propane'- propylene and butanes-butylenes, etc., from the vapors, and return them to main fractionator 22. The catalytically cracked gasoline obtained by means of line 27 has an octane number, on a debutanized basis, within the range of 84 to 94 research clear. The gasoline yields depend `on the processing conditions within the catalytic cracking unit, with a 31.5 API Mid-Continent gas oil, an 1l wt. percent alumina on silica cracking catalyst, a catalyst-to-oil ratio of 10, and a reactor temperature of 806 F., the yield of debutanized gasoline (l02-404 F. ASTM distillation range) is 42.6 vol. percent on feed, or 21,300 barrels per day on a debutanized basis, (By comparison, at 950 F. reactor temperature with the same charge stock and at the same convers-ion, the yield of debutanized gasoline is 37.2 vol. percent, or 18,600 barrels per day.) This gasoline has an hydrocarbon-type analysis of approximately 20 vol. percent C6 and heavier oils, 29% aromatics, 29% C6 and heavier parains and naphthenes, 8% pentenes and 14% pentanes. This gasoline, along with Ca-C., hydrocarbons if the gasoline is not subjected ltoV stabilization beforesplitting, is conducted to splitter fractionator 37.

In splitter fractionator 37--a tower operating -at atmospheric or slight subatmospheric pressure with a reboiler of 245 F. and an overhead temperature of 137 F.-the gasoline is carefully fractionated into a light catalytic gasoline component having an octane number higher than the debutanized gasolineand a heavy catalytic gasoline component with a lower octane number.v The eciency and economy of the over-all process depends to a large extent on ecient fractionation in splitter fractionator 37, and for this reason fractionator 37 is provided with 35 actual trays with a comparatively high reflux ratio of 2:1. The catalytic gasoline is thus split as closely and as accurately as possible at about 200 F. so that aliphatics (parans, olens, and naphthenes) in the light catalytic gasoline overhead boil below about the boiling point of n-heptane, while the heavy catalytically cracked gasoline bottoms contain aliphatics including n-heptane and higher boiling materials. In general, the fractionation is conducted so that the light gasoline contains less than about 10, and preferably less than about 2%, of total hydrocarbons boiling above 200 F. by true boiling point (TBP) distillation and the heavy catalytic gasoline bottoms contain less than l5, and preferably less than about 5 vol. percent, of hydrocarbons boiling below 200 F.

The economic importance of careful fractionation in splitter fractionator 37 cannot be overemphasized. While virtually all aromatic compounds within the gasoline boiling range have high octane numbers, irrespective of their identity, the octane numbers of aliphatics boiling below 200 F. are considerably higher than those boiling above this temperature. This may be shown with reference to Table I- below reporting computed octane numbers of three fractions distilled from a catalytic gasoline whose analysis is reported in the literature (Melpolder,l Ind` Eng. Chem., 44, 1142-46, 1952).

ases, eos

e Table I WEIGHTED AVERAGE RESEARCH OCTANE NUMBERS IBF-200 F. 20C-300 F. 30D-400 F. Fraction Fraction Fraction Paramus.. 80. 2 32. 4 33 Cycloparaffins 86. 3 63. 3 44 Olens 96. 2 55. 1 Cycloolefins 97. Y 98. 0 89 Aromatics 100. 0 118.0 113. 7

VOLUME PERCENT IN FRACTION Paratlins 14. 95 3. 27 6. 36 2.04 2. 71 5. 47 fins 25. 61 5. 06 3. 20 1.17 4. 69 1.47 Aromatics 0. 21 7. 73 16. 00

4moreover the expensive hydroforming operation would 'be applied to materials which already have a sufficiently Ahigh octane number.

The volume of heavy catalytic gasoline withdrawn from splitter fractionator 37 via line 39 (the light gasoline ,is obtained through line 38) is approximately 55% or 11,700 barrels per day based on the charge to the splitter. (At the conventional 950 F. operation the yield is 50% or 9,300 garrels per day.) This material is then conducted to olefin hydrogenation unit comprising charge pump 40, preheater 42, reactor 44, product cooler 46,

Vproduct separator 47, and a recycle hydrogen system including compressor 50. In the hydrogenation facilities, a sulfur insensitive catalyst which is selective for monoolefin and diolen hydrogenation without substantially saturated aromatic compounds is employed for the purpose of hydrogenating thermally unstable olefins which would form coke in the subsequent catalytic reforming zone. Suitable catalysts include platinum extended on alumina, which is preferred, molybdenum on alumina, cobalt Vmolybdate on alumina, nickel-tungsten sulfide, and other catalysts Well known in the art. The catalyst may be in the form of pills, pellets, balls, extrudates or the like, having a major dimension ranging from 1/32 to 1/2 inch or more. Since relatively low temperatures and high pressures are conducive to olefin hydrogenaton, conditions` obtaining in reactor 44 may be a temperature of from 300 F. .to about 800 F., preferably from about 400 F. to about 500 F., for platinum alumina catalysts, and the pressure may be from about 100 p.s.i.g. to about 1,000 p.s.i.g. or more. Excess hydrogen is employed in the system, and this may be used in a quantity at least ,stoichiometrically equivalent to the unsaturation of the heavy catalytic gasoline, and may be from about 100 to 1,000 standard cubic feet of hydrogen per barrel of liquid. To describe the olefin hydrogenation system in detail, the heavy catalytic gasoline obtained from splitter fractionator 37 through line 39 is pumped via charge pump 40 to juncture 41 where a stream of hydrogen-containing gas supplied through line 50, is commingled therewith land passed into preheater furnace 42 and thereafter sent by way of transfer line 43 to the top of hydrogenation reactor 44 containing a fixed bed of platinum alumina n hydrogen catalyst. The processing conditions obtaining in hydrogenation reactor 44 are typically an average .temperature of 500 F., an average pressure of 700 p.s.i.g., and a weight hourly space velocity of ve pounds of gasoline per hour per pound of catalyst; the catalyst comprises 0.3% platinum on alumina. After hydrogenaton, the reaction mixture is Withdrawn from reactor 44 and conducted through line 45 to product cooler 46 and thence sent to receiver-separator 47. In receiver-separator 47, the liquid hydrocarbon is separated and released via line 52 to acid products removal facilities to be described hereinafter, while hydrogen containing gas is taken by way of line 48 and recycled to the process by way of line 50 by compressor 49. A portion of the recycle gas is vented through line 51.

A major ancillary purpose of the olefin hydrogenation facilities is to hydrogenate nitrogen, sulfur, and chlorine-containing compounds in the heavy gasoline to form ammonium, hydrogen sulfide and hydrogen chloride. These materials are deleterious to platinum alumina reforming catalysts which are used at temperatures somewhat higher than employed for olefin hydrogenaton, and it is desirable to remove these materials prior to hydroforming. A preferred method of removal is shown in the attached ow sheet and comprises stripper 53 which contains a plurality of fractionation plates or trays; au ascending gas stream provided either by boiling a portion of the hydrogenated gasoline or by bubbling hydrogen-containing gas obtained from the catalytic reforming zone eectively strips gaseous ammonia, hydrogen sulfide and hydrogen chloride from the hydrocarbon and is re leased through vent 54.

The hydrogenated and preferably stripped heavy catalytic gasoline is then transmitted via line 55 to a catalytic hydroforming zone. Here the hydrocarbon is mixed with hydrogen-containing gas and contacted with a hydroforming catalyst to convert a substantial portion of the naphthenic and parainic constituents of the gasoline to aromatic compounds. As obtained from the olefin hydrogenation zone, the gasoline has quite a low octane number, generally in the range of about 70-80 research clear, due to the conversion of relatively high octane number olefins and cycloolefins to the corresponding lower octane number parafhns and naphthenes, and would be entirely unsuitable as a constituent in premium gasolines. However, by hydroforming over suitable catalysts--particularly platinum-alumina catalysts-these parafiins and naphthenes are converted to exceedingly high octane number aromatic compounds. Hydroforming catalysts are generally well known in the art, and are preferably 0.1-10% platinum extended on alumina; platinum-silicaalumina, palladium-alumina, alkalized chromina-alumina or the like. Hydroforming may be conducted nonregeneratively at pressures in the range of about 20G-1,000 p.s.i.g. or, with occasional or frequent regeneration, at pressures of 0-300 p.s.i.g., the average catalyst temperatures in each -case being Within the range of about 850- 1,050 F. Space velocities of from about 1.0 to 10.0, preferably from 2.0 to 5.0, e.g., 3.0, and hydrogen recycle rates of from 3,000 to 10,000 standard. cubic feet per barrel may be employed.

To follow the ow through the catalytic hydroforming zone, heavy catalytic gasoline supplied through line 55 may be blended with virgin naphtha obtained through valved line 56, and pumped by means of charge pump 58 through line 57 and thence to line 59, Where the liquid is mixed with recycle hydrogen-containing gas from line 78. The stream is then passed into preheater 60 which elevates the stream temperature to, say, 900 F., and the heated stream is then conducted through line 61 to the first stage reactor, reactor 62, containing platinumalumina catalyst. The predominant reaction in reactor 62 is dehydrogenation of naphthenes to aromatics, and

escasos the temperature is elevated to 925 F., and the stream transmitted through transfer line 65 to second stage reactor 66. In the latter reactor, additional naphthene dehydrogenation takes place, and substantial paraffin dehydrocyclization occurs; both of these reactions are endothermic land the reactor eluent is thereupon Withdrawn through line 67, reheated in reheater 68 and conducted through line 69 at a temperature of Vabout 950 F. At this temperature it is processed through reactor 70, also containing platinum-alumina catalyst, and sent through line 71 to cooler 72, and thence through line 73 to product separator 74. In separator 74, the liquid constituents are withdrawn through line 79, while the hydrogencontaining gas is collected by means of line 75 and recycled to the process by compressor 76 and line 78. Since hydroforming results in the production of net excess hydrogen, at least a portion of this is conuducted by way of line 77 to the olefin hydrogenation zone wherein the hydrogen requirements for olefin saturation are satisfied.

As a highly desirable embodiment of the catalytic hydroforming process, hydroforming is conducted at relatively low pressures on the order of 200-400 p.s.i.g.; under these conditions and at temperatures above about 900 F. the catalyst tends to become deactivated rapidly and must be regenerated by an interval ranging of from a few hours to a few days. To effect regeneration a socalled swing reactor 86 is provided, with transfer lines 83, 85 and 87 and reheater furnace 84. By means of suitable valves 83a and 87a, and corresponding valves and manifold lines connected to reactors 62, 66 and 70, any of the latter reactors may be temporarily Withdrawn from service, replaced with swing reactor 86 and regenerated by means of an oxygen-containing gas. The method of manifolding a `swing reactor and of regenerating la deactivated catalyst are fully described in U.S. Patent Number 2,773,014.

vHydroformed gasoline or reformate which is obtained from hydroforming zone separator 74 by means of line 79 is sent to stabilizer 80, a fractionating tower which removes light ends from the high octane gasoline via line 81. A portion of these light ends may be recycled by line 81a to the catalytic cracking zone for further conversion. By suitable control over the stabilizer operating pressure and temperature, a gasoline having any desired vapor pressure may be obtained at line 82.

A summary of processing conditions, yields, and product inspections in the present example is presented below.

Table Il SUMMARY oF EXPERIMENTAL DATA Reactor temperature, F. 806 Reactor pressure, lb./sq. in gage 5.0 Conversion, 100-vol. percent cycle oil 54.5 Dry gas, C3 and lighter1 3.5 Butylene fraction2 3.7 Butane fraction 2 7.4 C5 fraction 2 8.8 Depentanized gasoline2 33.8 Debutanized gasoline2 42.6 Total cycle oil2 45.4 Catalyst deposit 1 6.7

Gas yields, C3 and lighter, cu. ft./bbl./charge:

Hydrogen 45 Methane 6 Ethylene 8 Ethane 6 Propylene 59 Propane 22 Total 146 .1'Wt. percent of charge. Vol. percent of charge.

iC4H8 in total C4H8, vol. percent 14.0 iC4H10 in total CHm, vol. percent 90.4 C5H10 in total C5 fraction, vol. percent 24.0

Debutanized catalytic gasoline:

Gravity, API 59.8 Reid vapor pressure 6'. 1 Vol. percent C4 0.2 Total sulfur, wt. percent 0.10 Molecular weight3 99 Octane ratings- F-l clear 88.6 F-l-l-3 cc. TEL/gal. 95.2 F-2 clear 79.0 F2+3 cc. TEL/gat 84.3

ASTM distillation, 'Ff- IBP 102 10% 136 50% 219 358 aDerived from gravity and boiling point, Ind. Eng. Chem., 27, 1460 (1935).

The main product of the instant process, high octane gasoline obtained through line 82 has an octane number of research clear on a debutanized basis and is produced in a yield of 13,700 barrels per day or 27.4 vol. percent of the gas oil feed. It will be recalled that the catalytic cracking temperature is 806 F. and the heavy catalytically cracked gasoline before hydrogenation and hydroforming has an octane number research clear of 87.6. On the other hand, were catalytic cracking to be conducted at 950 F. a higher octane number heavy catalytic gasoline, 94.8 O N., would be obtained, and on hydrogen-ation and hydroforming would yield only 13,390 barrels per day (24.8 vol. percent on gas oil) of 100 octane number product.

From the foregoing discussion and specific embodiment, it is manifest that the process of the instant invention provides an extremely attractive yet economical method for producing high octane number gasoline blending stocks in large volume. By conducting a catalytic cracking of gas oil at moderate conditions to produce a debutanized gasoline of 84-94 octane number, fraction-ally distilling the cracked products to obtain a heavy catalytically cracked ,fraction boiling chiey Within the range of about 200 F. to about 400 F., hydrogenating olelins in this fraction and thereafter hydroforming the hydrogenated gasoline, extremely high yields of 98-110 octane number blending stocks are obtained. These stocks may be used either as an exceedingly high octane number unleaded premium or may be blended rwith other fnactions including the light catalytically cracked gasoline or extraneous virgin light gasolines, motor alkylatc, or polymer gasoline to provide premium blends.

Having described the invention, what is claimed is:

1. A process for producing high octane motor fuel blending stocks which comprises: catalytically cracking a gas oil in the presence of a uidized particulate solid cracking catalyst and at a temperature within the range of about 750 F. to about 900 F. to produce a debutanized catalytically cracked gasoline having an octane number within the range of about 84-94 research clear; separating a catalytically cracked gasoline from the reaction products; splitting said catalytically cracked gasoline to obtain a light high octane blending stock fraction and a heavy fraction boiling chiey With-in the range of 200 400 F.; recovering said light -high octane blending stock fraction; subjecting said heavy inaction to catalytic 'hydrogenation in the presence of hydrogen gas and a catalyst effective to selectively hydrogenate olenic constituents; subjecting the hydrogenated heavy fraction to cata-r lytic hydroforming in the presence of hydrogen `gas and a platinum hydroforming catalyst at a temperature of dan from about 800 F. to 1000 F., a pressure of from about 10 p.s.i.g. to about 1000 p.s.i.g., a space velocity of from about 1.0 to 10.0, und a hydrogen recycle rate of from 3,000 to 10,000 standard cubic feet per barrel to produce a debutanized reformate having an octane number within the range of about 98-110 research clear, at least a portion of the excess hydrogen from the hydroforming step being supplied to the olefin hydrogenation step; and recovering the reformate fraction as a second high octane blending stock fraction.

2. Process of claim 1 in which the cracking catalyst comprises uidized particulate silica-alumina.

3. Process of claim 1 in which the olen hydrogenation catalyst comprises platinum on alumina, and the olefin hydrogenation is effeoted at a temperature of from about 300 F. to Vabout 600 F. and a pressure of from about 100 p.s.i.g. to about 1,000 p.s.i.g.

4. Process of claim 1 in which the :olefin hydrogenation catalyst comprises cobalt molybda'te on alumina, and

the olefin hydrogenation is eiected at a temperature of from about 300 F. to about 600 F. and a pressure of from about 100 p.s.i.g. to about 1,000 p.s.i.g.

5. Process of claim 1 in which the hydroforming catalyst comprises platinum on alumina, and the hydroforming is eiected at a temperature of from about 800 F. to 1,000 F. and a pressure of from about 10 p.s.i.g. to about 1,000 p.s.i.g.

References Cited in the le of this patent UNITED STATES PATENTS 2,384,339 Read Sept. 4, 1945 2,417,308 Lee Mar. 11, 1947 2,442,276 Nelson et al. May 25, 1948 2,444,131 Delattre-Seguy June 29, 1948 2,740,751 Haensel Apr. 3, 1956 FOREIGN PATENTS 757,365 Great Britain Sept. 19, 1956 UNITED STATES PATENT OFFICE CERTIFICATION OF CORRECTION Parere No@ 2,968,605 Jenna-@y 1m wel Eugene S Coddau Jr et er,

It is hereby certfed'that error appears in the above numbered patent requiring correction and that the Said Lettere Patent should read as corrected below.

Column 5, line 40, for "garrels" read barrels mm column o, line 5l, for "'chromna" read chromla nm; column 7, line 17, for "eonuducted" read conducted column 8,

line 2, or "CHlO" read n C4Hl0 n Signed and sealed this 20th day of June IQl.,

(SEAL) Attest:

ERNEST W. SWIDER DAVID L.. LADD Attesting Officer Commissioner of Patents 

1. A PROCESS FOR PRODUCING HIGH OCTANE MOTOR FUEL BLENDING STOCKS WHICH COMPRISES: CATALYTICALLY CRACKING A GAS OIL IN THE PRESENCE OF A FLUIDIZED PARTICULATE SOLID CRACKING CATALYST AND AT A TEMPERATURE WITHIN THE RANGE OF ABOUT 750*F. TO ABOUT 900*F TO PRODUCE A DEBUTANIZED CATALYTICALLY CRACKED GASOLINE HAVING AN OCTANE NUMBER WITHIN THE RANGE OF ABOUT 84-94 RESEARCH CLEAR, SEPARATING A CATALYTICALLY CRACKED GASOLINE FROM THE REACTION PRODUCTS, SPLITTING SAID CATALYTICALLY CRACKED GASOLINE TO OBTAIN A LIGHT HIGH OCTANE BLENDING STOCK FRACTION AND A HEAVY FRACTION BOILING CHIEFLY WITHIN THE RANGE OF 200400*F., RECOVERING SAID LIGHT HIGH OCTANE BLENDING STOCK FRACTION, SUBJECTING SAID HEAVY FRACTION TO CATALYTIC HYDROGENATION IN THE PRESENCE OF HYDROGEN GAS AND A CATALYST EFFECTIVE TO SELECTIVELY HYDROGENATE OLEFINIC CONSTITUENTS, SUBJECTING THE HYDROGENATED HEAVY FRACTION TO CATALYTIC HYDROFORMING IN THE PRESENCE OF HYDROGEN GAS AND A PLATINUM HYDROFORMING CATALYST AT A TEMPERATURE OF FROM ABOUT 800*F. TO 1000*F., A PRESSURE OF FROM ABOUT 10 P.S.I.G TO ABOUT 1000 P.S.I.G., A SPACE VELOCITY OF FROM ABOUT 1.0 TO 10.0, AND A HYDROGEN RECYCLE RATE OF FROM 3,000 TO 10,000 STANDARD CUBIC FEET PER BARREL TO PRODUCE A DEBUTANIZED REFORMATE HAVING AN OCTANE NUMBER WITHIN THE RANGE OF ABOUT 98-110 RESEARCH CLEAR, AT LEAST A PORTION OF THE EXCESS HYDROGEN FROM THE HYDROFORMING STEP BEING SUPPLIED TO THE OLEFIN HYDROGENATION STEP, AND RECOVERING THE REFORMATE FRACTION AS A SECOND HIGH OCTANE BLENDING STOCK FRACTION. 